Conversion and desulfurization of hydrocarbonaceous black oils

ABSTRACT

A PROCESS FOR CONVERTING SULFUROUS, HYDROCARBONACEOUS BLACK OILS INTO LOWER-BOILING, NORMALLY LIQUID HYDROCARBON PRODUCTS OF REDUCED SULFUR CONTENT. THE PROCESS INVOLVES THE INTEGRATION OF HYDROGENATIVE CRACKING AND FIXEDBED CATALYTIC DESULFURIZATION, AND IS ESPECIALLY APPLICABLE TO THOSE HYDROCARBON CHARGE STOCKS CONTAINING LESS THAN 150 P.P.M. OF METALLIC CONTAIMINANTS. THE CHARGE STOCK IS INITIALLY SUBJECTED TO FIXED-BED CATALYTIC HYDROGENATION AND DESULFURIZATION. FOLLOWING SEPARATION OF THE CATALYTIC REACTION ZONE PRODUCT EFFLUENT, A HIGH-BOILING CONCENTRATE IS THERMALLY-CRACKED IN THE PRESENCE OF DISSOLVED HYDROGEN.   D R A W I N G

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MQQMQIMQ QQWU QN bPGl www@ QQ /N VE /V TOR" Fran/r Sto/fa A TTOY/VEYS I'United States Patent O 3,594,309 CONVERSION AND DESULFURIZATION F HYDROCARBONACEOUS BLACK OILS Frank Stolfa, Park Ridge, Ill., assignor to Universal Oil Products Company, Des Plaines, Ill. Filed Oct. 28, 1968, Ser. No. 771,248 Int. Cl. Cg 37/04 U.S. Cl. 208-89 4 Claims ABSTRACT OF THE DISCLOSURE A process for converting sulfurous, hydrocarbonaceous black oils into lower-boiling, normally liquid hydrocarbon products of reduced sulfur content. The process involves the integration of hydrogenative cracking and fixedbed catalytic desulfurization, and is especially applicable to those hydrocarbon charge stocks containing less than 150 ppm. of metallic contaminants. The charge stock is initially subjected to fixed-bed catalytic hydrogenation and desulfurization. Following separation of the catalytic reaction zone product eluent, a high-boiling concentrate is thermally-cracked in the presence of dissolved hydrogen.

APPLICABILITY OF INVENTION The process described herein is adaptable to the desulfurization of petroleum crude oil residuals having relatively low metals content-ie. containing less than about 150 p.p.m. of total metals. More specifically, the present invention is directed toward a combination process for converting and reducing the sulfur concentration of hydrocarbonaceous charge stocks commonly referred to in the art as black oils.

Petroleum crude oils, and particularly the heavy residuals extracted from tar sands, topped or reduced crudes, vacuum residuals, contain high molecular Weight sulfurous compounds in exceedingly large quantities, nitrogenous compounds, asphaltic material, insoluble in light hydrocarbons such as pentane and heptane, and high molecular weight organo-metallic complexes. With respect to the metallic complexes, containing nickel and vanadium as the principal metallic components, the various black oil charge stocks can be classified as (l) high metals residuals or (2) low metals residuals. The present invention is primarily directed to the processing of those hydrocarbonaceous black oils having a low metals content of less than about 150 p.p.m., computed as if existing in the elemental state. A Iblack oil is generally characterized as a heavy lcarbonaceous material of which more than about 10.0% by volume has a normal boiling point above a temperature of 1050 F. (referred to as nondistillables). Such material generally has a gravity less than about 20.0 API and sulfur concentrations greater than about 2.0% by weight. Conradson carbon residue factors exceed 1.0% by Weight, and a great proportion of black oils indicate a Conradson carbon residue factor above 10.0.

Exemplary of those hydrocarbonaceous black oils, to the conversion and desulfurization of which the present invention is directed, include a crude tower bottoms product having a gravity of about 14.3 API, and contaminated by the presence of about 3.0% by weight of sulfur, 3830 p pm. of total nitrogen, S5 ppm. of total metals and about 11.0% by weight of asphaltenes. Another typical charge stock is a vacuum column bottoms product derived from a Middle-East crude oil. This vacuum bottoms product has a gravity of `60 API, an average molecular weight of about 620, an ASTM 20.0% volumetric distillation temperature of about 1035 F., and contains about 4000 p.p.m. of nitrogen, 5.5% by ice weight of sulfur, p.p.m. of vanadium and nickel, and 6.0% by weight of heptane-insoluble aspheltenes. The present invention affords the conversion of such material into lower-boiling, normally liquid hydrocarbon products, and further converts a considerable quantity of nondistillables. Additionally, the normally liquid product of the process has been substantially reduced in sulfur content-ie. less than about 1.0% by weight of sulfur.

The principal diiculty, heretofore encountered in the processing of hydrocarbonaceous black oils, resides in the lack of a significant degree of sulfur stability of catalytic composites when the charge stock to be processed is characterized by the presence of large quantities of asphaltic material. This difficulty arises primarily as a consequence of the necessity for effecting the process at an operating severity level such that non-distillable conversion simultaneously takes place while sulfurous compounds are being converted into hydrogen sulfide and hydrocarbons. The asphaltic material, dispersed within the charge stock, has the tendency to flocculate and polymerize, whereby the conversion thereof to more valuable oil-soluble products is virtually precluded. Furthermore, the sulfur-containing polymerized asphaltic complexes become deposited upon the catalytic composite, steadily increasing the rate at which the catalytic composite becomes deactivated. These difficulties are further compounded with respect to those charge stocks characterized by a high metals content. Since these charge stocks contain metals in amounts as high as 700 ppm., the catalyst deactivation rate is accelerated to the extent that processing to produce lower-boiling hydrocarbon products is not economically feasible.

The present invention is founded on recognition of the fact that an acceptable degree of desulfurization of low metals-containing black oils is possible at relatively mild operating severities which favor extended catalyst life without effecting a significant degree of asphaltene polymerization. In order that the process becomes economically attractive from the standpoint of producing lower-boiling hydrocarbon products, an essential feature of my invention resides in the subsequent processing of the liquid product efiiuent from the fixed-bed catalytic reaction zone. Therefore, as hereinafter set forth in greater detail, the reaction zone effluent is separated at a temperature of from about 100 F. to about 800 F., and at substantially the same pressure as imposed upon the catalytic reaction zone, in order to provide a principally liquid phase, at least a portion of which is subsequently subjected to a noncatalytic, thermal cracking reaction zone, or coil.

OBJECTS AND EMBODIMENTS A principal object of my invention is to provide an economical process for effecting the desulfurization and hydrogenative conversion of low metal-containing black oils. A corollary objective is to extend the period of acceptable, economical catalyst life while desulfurizing and converting hydrocarbonaceous black oils containing less than about p.p.m. of total metals.

Another object is to convert heavy hydrocarbon charge stocks, a significant amount of which exhibits a boiling range above a temperature of 1050 F., into lower-boiling distillable hydrocarbons having a sulfur concentration less than about 1.0% by weight.

Therefore, in one embodiment, my invention relates to a process for the con-version of a sulfurous, hydrocarbonaceous charge stock, of which at least about 10.0% boils above a temperature of about 1050 EF., into lower-boiling hydrocarbon products, which process comprises the steps of: (a) heating said charge stock to a temperature of from 500 F. to about 750 F., reacting said charge stock with hydrogen in a catalytic reaction zone, and in contact therein with a catalytic composite at a pressure greater than about 1000 p.s.i.g.; (ib) separating the resulting reaction zone eflluent, in a irst separation zone, at substantially the same pressure imposed upon iirst catalytic reaction zone, to provide a first vapor phase and a first liquid phase; (c) separating said first vapor phase, in a second separation zone, at substantially the same pressure imposed upon said iirst separation zone, to provide a second liquid phase and a second vapor phase rich in hydrogen; (d) cracking at least a portion of said rst liquid phase in a non-catalytic second reaction zone; (e) introducing the resulting cracked product effluent into a fractionation zone, and also introducing into said fractionation zone said second liquid phase at a locus above said cracked product eiuent; (f) withdrawing a hydrocarbon stream boiling substantially above a temperature of about 650 F. from said fractionation zone, and introducing said stream into a third separation zone; and, (g) separating said stream at a reduced pressure of from subatmospheric to about 50 p.s.i.g., to provide a third liquid phase boiling below a temperature of about 1050 F. and a residuum concentrate.

Other embodiments of my invention, as hereinafter set forth in greater detail, reside primarily in preferred ranges of process variables and in various processing techniques. For example, the total charge to the iirst iixed-bed catalytic reaction zone, consisting primarily of fresh charge stock, a recycled portion of the rst liquid phase, a recycled hydrogen-rich vapor phase and make-up hydrogen, required to supplant that which is consumed within the overall process and to maintain pressure, is heated to a temperature within the preferred range of from about 650 F. to about 750 F The precise temperature is controlled within the aforesaid range by monitoring the temperature of the reaction zone product effluent. Since the principal reactions being effected are highly exothermic, a temperature rise is experienced as the charge stock and hydrogen passes through the catalyst bed. In many instances, some temperature control is afforded through the use of a quench stream. Economically acceptable catalyst life is achieved when the maximum catalyst temperature, which is virtually the same as that of the product eiiiuent, is maintained at a level below about 800 F. In another embodiment, the iirst reaction zone effluent being introduced into the irst separation zone, is at a temeprature of from about 700 F. to about 800 F. in order that the portion of the first liquid phase being subjected to the subsequent thermal cracking reaction zone, or coil, contains from about 10.0 mol. percent to about 40.0% of dissolved hydrogen. Other objects and embodiments of my invention will be evident from the following, more detailed description of the process encompassed thereby.

SUMMARY OF INVENTION As hereinbefore set forth, the principal function of the present invention resides in the production of maximum quantities of distillable hydrocarbons which have been reduced in sulfur concentration. Through the utilization of the present process, this is accomplished in a highly economical fashion while avoiding the difficulties and pitfalls of currently-practiced processing techniques. Paramount is the extension of the period of time during which the fixed-bed catalytic composite functions in an acceptable manner. With respect to the processing of high metals black oils, being those containing more than about '150 p.p.m. of total metals, it has been found that a successful operation involves initially hydrovisbreaking the fresh hydrocarbon charge stock in the presence of limited quantities of hydrogen. While both technical and economical justification exists to support this processing technique, particularly respecting the attainable catalyst life experienced in the fixed-bed reaction zone, there is incurred a yield loss with respect to that quantity of the original nondistillable asphaltics which are not converted. This yield loss stems principally from the fact that thermal cracking does not achieve the conversion of all the convertible asphaltics within the charge stock, the unconverted portion of which is removed as a heavy residuum prior to subjecting the remainder of the thermally-cracked product effluent to additional conversion in a fixed-bed catalytic reaction zone. If the as-received high metals charge stock were processed initially in the fixed-bed catalytic reaction zone, the presence of the exceedingly high concentration of metals, in an environment conducive to effecting acceptable desulfurization, results in extremely rapid catalyst deactivation. In accordance with the present process, primarily applicable to those charge stocks of low metals content, the residual charge stock is catalytically desulfurized, and at least partially converted, at relatively mild hydrogenation severities which favor extended catalyst life. The catalytically converted product effluent is separated into a principally vaporous phase and a principally liquid phase, at least a portion of the latter being utilized as the charge to a non-catalytic thermal cracking reaction zone.

In a preferred embodiment, the total charge to the xed-bed catalytic reaction zone includes the fresh hydrocarbon charge stock, a recycled hydrogen-rich gaseous phase, make-up hydrogen and a recycled diluent, the source of the latter being hereinafter set forth. 'This mixture is raised to a temperature of from about 500 F. to about 750 F., and preferably from about 650 F. to about 750 F. as measured at the inlet to the catalyst bed. In order to preserve catalyst stability, the inlet temperature is controlled at a level such that the temperature of the reaction product effluent, or the maximum catalyst bed temperature does not exceed about 800 F. The reaction zone will be maintained under an imposed pressure of from about 1000 to about 4000 p.s.i.g., and the hydrocarbon charge stock will contact the catalytic composite at a liquid hourly space velocity of from about 0.5 to about 10.0, based upon the fresh hydrocarbon charge stock, exclusive of any recycled diluent. The hydrogen concentration will be in the range of from about 5000 to about 50,000 standard cubic feet per barrel, while the combined feed ratio, defined as total volumes of liquid charge per volume of fresh hydrocarbon charge, is in the range of from about 1.1:1 to about 3.5:1.

The catalytic composite disposed within the fixed-bed catalytic reaction, or conversion zone, can be characterized as containing a metallic component having hydrogenation activity, which component is combined with a suitable refractory inorganic oxide carrier material of either synthetic, or natural origin. The precise composition and method of manufacturing the carrier material is not considered essential to the present invention, although a siliceous carrier, such as 88.0% by weight of alumina and 12.0% by weight of silica, or 63.0% by weight of alumina and 37.0% by weight of silica, or 68.0% by weight of alumina, 10.0% by weight of silica and 22.0% by weight of boron phosphate are generally preferred. Suitable metallic components having hydrogenation activity are those selected from the `group consisting of the metals of Group VI-B and VIII of the Periodic Table, as set forth in the Periodic Table of the Elements, E. H. Sargent & Company, 1964. Thus, the catalytic composites may comprise one or more metallic components from the group of molybdenum, tungsten, chromium, iron, cobalt, nickel, platinum, iridium, osmium, rhodium, ruthenium, and mixtures thereof. The concentration of the catalytically active metallic component, or components, is primarily dependent upon a particular metal as well as the physical and/ or chemical characteristics of the charge stock. For example, the metallic components of Group VI-B are generally present in an amount within the range of from about 1.0% to about 20.0% by weight, the iron-group metals in an amount within the range of about 0.2% to about 10.0% by weight, whereas the noble metals of Group VIII are preferably present in an amount within the range of from about 0.1% to about 5.0% by weight, all of which are calculated as if these components existed within the catalytic composite in the elemental state.

Before further summarizing my invention, several definitions are believed necessary in order that a greater understanding of the invention is afforded. In the present specification, and in the appended claims, the terms principally vaporous and principally liquid, are intended to describe a particular stream, the major proportion of the components of which are either normally gaseous, or normally liquid at standard conditions. Similarly, the phrase pressure substantially the same as is intended to connote that the pressure under which a succeeding vessel is maintained, is the same as the previous vessel, allowing only for the pressure drop experienced as a result of the flow of fluids through the system. For example, Where the catalytic first reaction zone is maintained at a pressure of about 2900 p.s.i.g., the yfirst separation zone, or hot separator will function at about 2780 p..i.g. The phrase temperature substantially the same as is employed to in dicate that the only reduction in temperature stems from the normally experienced loss due to the -flow of material from one piece of equipment to another, or from the conversion of sensible heat to latent heat by flashing in which a pressure drop occurs. It must further be stated that the distillable portion of the charge stock comprises those hydrocarbons which can be distilled at a temperature lbelow about 1050 F. As a practical matter, since the charge stock contains asphaltics, this temperature becomes limiting in order to prevent cracking. However, with respect to the normally liquid product, from `which the asphaltic residuum has been removed, analyses have indicated end boiling points as high at 1100 F. to 1150 F.

The total product effluent from the first catalytic reaction zone, at a maximum temperature of about 800 F., and preferably at a maximum temperature of about 775 F., is passed into a first separation zone hereinafter referred to as the hot separator. The principal function served by the hot separator is to separate the mixed-phase product effluent into a principally Vapor phase rich in hydrogen and a principally liquid phase containing from about 10.0 mol percent to about 40.0 mol percent of dissolved hydrogen. In a preferred embodiment, the total reaction product effluent is utilized as a heat-exchange medium in order to lower the temperature thereof to a level in the range of from about 700 F. to about 800 F., and preferably below a level of 750 F. The principally vaporous phase from the hot separator is introduced into a second separation zone hereinafter referred to as the cold separator. The cold separator, operating at substantially the same pressure as the hot separator, but at a significantly lower temperature in the range of about 60 F. to about 140 F., serves to concentrate the hydrogen in a second principally vaporous phase. This hydrogenrich vaporous phase, comprising about 80.0 mol percent hydrogen, and only about 2.2 mol percent propane and heavier hydrocarbons, is made available for use as a recycle stream to be combined with the fresh black oil charge stock. Butanes and heavier hydrocarbons are condensed in the cold separator, and removed therefrom as a second principally liquid phase.

The liquid phase from the hot separator is in part recycled to combine with the fresh hydrocarbon charge stock to serve as a diluent for the heavier constituents thereof. The quantity of the liquid phase diverted in this manner is such that the combined feed ratio to the catalytic reaction Zone, being defined as total Volumes of liquid charge per volume of fresh liquid charge, is within the range of from about 1.1:1 to about 3.5: 1. The remaining portion of the principally liquid phase from the hot separator is introduced into a thermal cracking reaction zone, or coil, at a reduced pressure in the range of from about 200 p.s.i.g. to about 500 p.s.i.g. and at a temperature of from about 700 to about 750 F.

The thermally-cracked product effluent is introduced into a fractionation zone, or distillation column maintained at conditions of temperature and pressure such that light, normally gaseous hydrocarbons and gasoline boiling range hydrocarbons, having a nominal end boiling point of about 400 F., are removed as an overhead fraction, middle-distillate hydrocarbons boiling from about 400 F. to about 650 F. are removed as a side-cut fraction, and material boiling above about 650 F. are removed as a bottoms fraction. As hereinafter indicated in the description of the accompanying drawing, the second principally liquid phase from the cold separator is also introduced into the distillation column, and at a locus intermediate that from which the middle-distillates are withdrawn, and that at which the thermally-cracked product effluent is introduced. The 650 F. plus hydrocarbonaceous material, withdrawn from the bottom portion of the distillation column, is introduced into a vacuum flash column maintained at about 20 to about 60 mm. Hg absolute. The vacuum flash zone serves as the third separation zone, the principal function of which is the concentration and recovery of an asphaltic residuum substantially free from distillable hydrocarbons. In general, vacuum gas oil streams are recovered from the vacuum flash column as a separate light vacuum gas oil (LVGO) and a heavy vacuum -gas oil (HVGO), although in some instances, a medium vacuum gas oil is also recovered. In addition, in one preferred embodiment, a slop-wax stream, boiling within a temperature range of about 980 F. to about 1150 F., is withdrawn from the vacuum flash column, and at least a part recycled to combine with the fresh hydrocarbon charge stock in an amount of from 1.0% to 20.0%. When so recycled, the amount of the slop-wax is such that the combined feed ratio to the catalytic reaction zone, including the recycled portion of the first liquid phase from the hot separator, continues to be maintained in the range of from about 1.1:1 to about 3.5 :1, as hereinafter set forth. In another embodiment, a portion of the slop-wax and/or heavy vacuum gas oil, may be recycled to combine with the first liquid phase being introduced to the thermal cracking reaction zone. The amount so recycled is such that the combined feed ratio to the thermal cracking reaction coil is above about 1.2:l, and generally not higher than about 4.011.

The principal advantage, or benefit, attendant the use of my invention, resides in an extension of the period of acceptable catalyst life with respect to a fixed-bed catalytic reaction zone. This stems primarily from the fact that desulfurization, to a level less than about 1.0% by weight, is effected at a relatively low severity of operation with the result that the atmosphere within the reaction zone is not conducive to the formation of polymer products, containing sulfur, Otherwise resulting from the presence of hydrocarbon-insoluble asphaltenes. Another advantage resulting from the low severity of operation within the fixed-bed catalytic reaction zone, involves the dissolution of hydrogen in the normally liquid heavier portion of the product effluent, at least a portion of which is utilized as the charge to the thermal cracking reaction zone. Of further interest is the fact that the vacuum flash column is significantly reduced in size as a result of the fractionation zone being utilized to separate the thermal reaction zone product effluent to concentrate the material boiling above a temperature of about 650 F. This, as will be recognized by those having skill in the art of petroleum processing techniques, affords an added advantage with respect to the overall economics of the process. The introduction of the second principally liquid phase from the cold separator, into the distillation column at a point intermediate the withdrawal of the middledistillates and the introduction of the thermally-cracked product effluent, further insures that the bottoms stream contains a minimum quantity of material boiling below 650 F. Since the hydrocarbonaceous material boiling below about 650 F. is separately recovered in the distillation zone, there is effected a beneficial reduction in the size of the vacuum flash column.

7 DESCRIPTION OF DRAWING For the purpose of demonstrating the illustrated embodiment, the drawing will be described in connection with the conversion and desulfurization of a vacuum column bottoms product derived from a full boiling range crude oil. The vacuum bottoms product has a gravity of 11.7 API, an average molecular weight of about 600, an ASTM initial boiling point of about 860 F. with about 44.0% by volume being distillable at a temperature of 1050 F., contains 4600 p.p.m. of nitrogen, 1.92% by weight of sulfur, 105 p.p.m. of vanadium and nickel, has a Conradson carbon residue factor of 12.8% by weight and contains about 3.45% by weight of heptaneinsoluble asphaltenes.

In addition, the description will be directed toward a commercially-scaled unit having a capacity of about 4000 barrels per day. In the drawing, the embodiment is presented by means of a simplified flow diagram in which details such as pumps, instrumentation and controls, heatexchange, and heat-recovery circuits, valving, start-up lines and similar hardware have been omitted as nonessential to an understanding of the techniques involved. The utilization of such miscellaneous appurtenances, to modify the illustrated process flow, are well within the purview of those skilled in the art. Similarly, it is understood that the charge stock, stream compositions, operating conditions, design of fractionators, separators and the like are exemplary only, and may be varied widely without departure from the spirit of my invention, the scope of which is defined by the appended claims.

This vacuum column bottoms is intended for conver- -sion to maximum distillable hydrocarbons recoverable by ordinary distillation and commonly utilized fractionation facilities. The charge stock is processed in a fixed-bed catalytic conversion zone in admixture with about 10,000 standard cubic feet per barrel of hydrogen, based upon fresh feed, at an inlet catalyst bed temperature of about 685 F. and an inlet pressure of about 2650 p.s.i.g. The liquid hourly space velocity, based upon fresh feed only, is 0.5 and the combined feed ratio, with respect to total liquid feed, is about 2.0: 1. In this particular instance, the product streams are intended to be a 400 F. end point gasoline fraction, a 400 F. to 500 F. kerosene cut, a middle-distillate fraction boiling from 500 F. to 650 F., and a gas oil stream having an intial boiling point of about 650 F. and containing all the remaining distillable hydrocarbons in the product eflluent.

With reference now to the drawing, the vacuum column bottoms, in an amount of about 3,600 bbl/day, is introduced into the process by way of line 1, is admixed with about 2.0% by volume of a slop-Wax recycle from line 2, 3,600 bbl./day of a hot separator liquid stream in line 3 and a recycled hydrogen-rich gaseous stream in line 4 (about 10,000 s.c.f./bbl. of hydrogen), the mixture continuing through line 1 into heater 6. Not illustrated in the drawing is the technique whereby the total charge to heater 6 is first heated by way of conventional heatexchange with various hot effluent streams, 'and to a temperature of about 635 F. Heater 6 raises the temperature of the feed mixture to 685 F., the heated mixture passing through line 7 into fixed-bed catalytic reactor 8.

The reactor product effluent, in mixed phase in line 9, at a temperature of 785 F., is utilized as a heat-exchange medium, in order to lower the temperature to 750 F., and enters hot sepaartor 10 at a pressure of about 2600 p.s.i.g. A principally vaporous phase is withdrawn from hot separator 10 via line 11 and, after cooling by conventional means, is introduced into cold separator 12 at about 100 F. and a pressure of 2550 p.s.i.g. A principally liquid phase is withdrawn from separator 10, by way of line 15, and a portion thereof (3,600 bbL/day) is diverted through line 3 to combine with the fresh charge in line 1 and the slop-wax recycle (about 72 bbl./day) from line 2. The remainder continues via line into thermal reaction coil 16 at a pressure of about 250 p.s.i.g. and a temperature of about 750 F. A hydrogen-rich gaseous phase is removed from cold separator 12 by way of line 4 through the use of compressive means not illustrated in the drawing. After make-up hydrogen is introduced via line 5, the gaseous mixture continues through line 4 to be combined with the liquid feed mixture in line 1. A liquid phase from cold separator 12 is introduced into fractionator, or distillation column 14 by way of line 13.

By way of further illustrating the process, as thus far set forth, the following Tables I and -II present the component analyses of the streams resulting from the eparations effected in hot separator 10 and cold separator 12. The indicated analyses, in rnol percent, do not account for various recycle streams and/or quench streams. The analyses for the principally vaporous phase in line 11 and the liquid phase in line 15 are given in the following Table I.

TABLE I.-HOT SEPARATOR STREAM ANALYSES Line Number Component stream, mol. percent* Nitrogen Hydrogen Hydrogen sulfide. CrC3 l Hydrogen is 38.0 mol. percent. 2 About 17.1% residuum.

Approximate analyses for the liquid phase in line 13 and the vapor phase in line 4 are presented in Table II.

TABLE II.-COLD SEPARATOR STREAM ANALYSES Line Number 4 13 750 F.-plus 0. 5

The principally liquid phase from cold separator 12 is introduced into fractionator 14 by way of line 13, at a locus above that at which the thermally-cracked product euent is introduced thereto via line 17. The thermallycracked effluent is at a pressure of about p.s.i.g., and a temperature of 930 F. These will be changed, as will the temperature and pressure of the cold separator liquid in line 13 to the extent necessary to conform to the conditions imposed upon the distillation column in order to provide the desired product cuts. In the embodiment illustrated, which is a commonly desired system, fractionator 14 will function to provide an overhead fraction consisting of a minor amount of normally gaseous components and those normally liquid hydrocarbons in the gasoline boiling range-ie. boiling below about 400 F.- shown leaving via line 18. A kerosene fraction, boiling in the range of from 400 F. to 500 F. and a middledistillate fraction, boiling from about 500 F. to about 650 F., shown being withdrawn by way of lines 19 and 20, respectively. It should be noted that the locus of withdrawal of the middle-distillate fraction is above that at which the cold separator liquid is introduced. Since the ,highest end boiling point of the desired product streams rrom the fractionator is 650 F., in the illustration, the fractionator is controlled at conditions which maximize the concentration of 650 F.-plus material leaving via line 21. Therefore, the thermally-cracked product eluent in line 17 is quenched to a temperature of about 800 F.

The heavier material, at a temperature of about 800 F. is introduced into vacuum flash zone 22 functioning at about 30 mm. of Hg, absolute. The vacuum flash column serves to concentrate a residuum fraction, line 25, which may be blended with thc heavy gas oil (850 F. to about 980 F.), line 24, and at least a portion of the slop-wax (980 F.1150 F.) not being recycled through line 2. This mixture of slop-wax, residuurn and heavy vacuum gas oil is well-suited as a fuel oil since its sulfur content is well within the common specification of 1.0% by weight. A light vacuum gas oil is removed from vacuum flash zone 22 by way of line 23. The relatively minor quantity of hydrocarbonaceous material boiling below 650 F., is removed from the vacuum flash column by conventional jets which are not illustrated in the drawing.

The overall product yields, and distribution, are presented in the following Table III:

TABLE III.-YIELDS AND PRODUCT DISTRIBUTION Component: Yield Ammonia, wt. percent 0.15 Hydrogen sulfide, wt. percent 0.70 C'l-C, wt. percent 3.63 Butanes, vol. percent 2.31 Pentanes, vol. percent 1.47 Hexanes, vol. percent 2.29 C7- 400 F., vol. percent 11.19 400 F.-500 F., vol. percent 11.44 500 F.-650 F., vol. percent 21.18 650 F.-plus, vol. percent 37.02 Residuum 18.61

Of further interest is the fact that the C7-400 F. fraction has an API gravity of 46.7 and a sulfur content less than 0.1% by weight; that of the 400 F.-500 F. fraction is about 0.210% by weight with a gravity of 34.5 API; the 500 F. lto 650 F. fraction has a gravity of about 29.6 API, and contains about 0.2% sulfur; and, the `650" F.-plus portion has a concentration of 0.25% byweight of sulfur and a gravity of 19.0 API.

On the basis of the 3600 bbl./day of fresh charge stock, 86.9 vol. percent of the total C4,plus product is distillable, or 3130 bbL/day. In the case where the residuum, 18.61 vol. percent, is blended with the heavy vacuum gas oil to form a fuel oil, the total yield of useable normally liquid products becomes about 3800 bbl./ day. Significantly, in the scheme illustrated, this is accomplished with the consumption of hydrogen of only about 875 s.c.f./bb1. of fresh charge, or about 1.31% by weight.

I claim as my invention:

1. A process for the conversion of a sulfurous hydrocarbonaceous charge stock, containing less than about 150 p.p.m. metallic contaminants, and of which at least about 10.0% boils above a temperature of about 1050 F., into lower-boiling hydrocarbon products, which process comprises the steps of:

(a) heating said charge stock to a temperature of from 500 F. to about 750 F., reacting said charge stock with hydrogen in a catalytic reaction zone, in contact with a catalytic compositey and at a pressure greater than about 1000' p.s.i.g.;

(b) separating the resulting reaction zone effluent in a first separation zone, at substantially the same pressure imposed upon said first reaction zone and at a temperature from 700 F. to about 800 F., to provide a first vapor phase and a first liquid phase.

(c) separating said first vapor phase, in a second separation zone, at substantially the same pressure imposed upon said first separation Zone, to provide a second liquid phase and a second vapor phase rich in hydrogen;

(d) cracking at least a portion of said first liquid phase, without intermediate heating thereof, in a non-catalytic second reaction zone;

(e) introducing the resulting cracked product efliuent into a fractionation zone, and also introducing into said fractionation zone, said second liquid phase at a locus above said cracked product effluent;

(f) withdrawing a hydrocarbon stream boiling substantially completely above a temperature of about 650 F. from said fractionation zone, and introducing said stream into a third separation zone;

(g) separating said stream, at a reduced pressure of from subatmospheric to about 50 p.s.i.g., to provide a third liquid phase containing distillable hydrocarbons and a residuuni concentrate; and

(h) recycling at least a portion of said third liquid phase to combine with said charge stock.

2. The process of claim 1 further characterized in that at least a portion of said second vapor phase is recycled to combine with said charge stock.

3. The process of claim 1 further characterized in that at least a portion of said first liquid phase is recycled to combine with said charge stock, to provide a combined liquid feed ratio to said rst reaction zone in the range of from 1.1:1 to about 3.5:1.

4. The process of claim 1 further characterized in that at least a portion of said third liquid phase is recycled to combine with said first liquid phase to provide a combined feed ratio to said non-catalytic second reaction zone above about 1.211.

References Cited UNITED STATES PATENTS 1,932,174 10/1933 Gaus et al. 196-24 2,282,451 5/ 1942 Brooks 196-24 2,327,099 8/1943 Eastman 196-49 2,339,918 1/1944 Thomas 196-52 2,355,366 8/1944 Conn 196-24 3,409,538 ll/l968 `Gleim et al. 208-59 DELBERT E. GANTZ, Primary Examiner R. M. BRUSKIN, Assistant Examiner U.S. Cl. XR. 208-100, 102, 6l 

